Apparatus for preparing polyolefin products and methodology for using the same

ABSTRACT

Apparatus for olefin polymerization includes a plurality of shell and tube olefin polymerization reactors, each of which has an olefin polymerization reaction mixture inlet connection and a crude polyolefin product outlet connection. Each reactor is equipped with a recirculation system including a pump arranged to circulate a reaction mixture through the tube side of the reactor independently of the introduction of olefin polymerization reaction mixture into the reactor. The apparatus also includes an inlet reaction mixture distribution manifold and an outlet polymerization reaction mixture collection manifold interconnecting the reactors for operation in parallel. The apparatus also includes catalyst composition and catalyst modifier inlets for each reactor arranged such that a catalyst modifier to may be introduced into each reactor at a rate which is independent of the introduction of catalyst composition. The apparatus further incorporates a crude polyolefin product catalyst removal and wash system including a plurality of settler vessels, associated piping and an inlet for a catalyst killing agent. This crude polyolefin product catalyst removal and wash system operates to receive crude polyolefin product from the crude polyolefin product outlet and remove residual catalyst therefrom.

CROSS REFERENCE TO RELATED APPLICATIONS

This application is a divisional of and claims priority pursuant to 35U.S.C. § 120 from co-pending application Ser. No. 10/434,805, filed May9, 2003. The entirety of the disclosure of said prior application Ser.No. 10/434,805 is hereby specifically incorporated herein by thisspecific reference thereto.

BACKGROUND OF THE INVENTION

1. Field of the Invention

The present invention relates to liquid phase olefin polymerization, tothe preparation of polyolefin products and to apparatus useful in thepreparation of polyolefin products. In particular the present inventionrelates to apparatuses and equipment for the preparation of a variety ofpolyolefin products using a liquid phase polymerization process and tothe methodology used in the operation of such apparatuses and equipment.More particularly, the present invention relates to apparatus andmethodology which enhances the operation and control of polyolefinreactors.

2. Background of the Invention

Presently pending U.S. patent application Ser. No. 09/515,790 filed onFeb. 29, 2000 and entitled “Process For Producing High VinylidenePolyisobutylene” (hereinafter the '790 application) relates to liquidphase polymerization processes for preparing low molecular weight,highly reactive polyisobutylene. In accordance with the disclosure ofthe '790 application, a catalyst composition, which desirably maycomprise a complex of BF₃ and methanol, and a feedstock containingisobutylene, are each introduced into a reaction zone where the same areintimately admixed with residual reaction mixture so as to present anintimately intermixed reaction admixture in the reaction zone. Theintimately intermixed reaction admixture is maintained in its intimatelyintermixed condition and at a relatively constant temperature of atleast about 0° C. while the same is in the reaction zone, wherebyisobutylene therein is polymerized to form polyisobutylene (PIB) havinga high degree of terminal unsaturation. A crude product streamcomprising residual catalyst composition, unreacted isobutylene andpolyisobutylene is then withdrawn from the reaction zone. Theintroduction of feedstock into and the withdrawal of product stream fromthe reaction zone are each controlled such that the residence time ofthe isobutylene undergoing polymerization in the reaction zone is nogreater than about 4 minutes, whereby the product stream contains ahighly reactive polyisobutylene product. Preferably, the reaction zonemay be the tube side of a shell-and-tube exchanger in which a coolant iscirculated on the shell side. A recirculation loop may desirably beemployed to circulate the reaction admixture through the tube sidereaction zone at a linear velocity sufficient to establish and maintainan intimately intermixed condition in the admixture and remove heatgenerated by the exothermic polymerization reaction.

U.S. Pat. No. 6,525,149 issued on Feb. 25, 2003 and entitled “ProcessFor Preparing Polyolefin Products” (hereinafter the '149 patent) relatesto a novel liquid phase polymerization process for preparing apolyolefin product having preselected properties. The process of the'149 patent includes the steps of providing a liquid feedstock whichcontains an olefinic component and a catalyst composition which maycomprise a stable complex of BF₃ and a complexing agent. The feedstockmay comprise any one or more of a number of olefins, including branchedolefins such as isobutylene, C₃-C₁₅ linear alpha olefins and C₄-C₁₅reactive non-alpha olefins. The feedstock and the catalyst compositionmay desirably be introduced into a residual reaction mixturerecirculating in a loop reactor reaction zone provided on the tube sideof a shell and tube heat exchanger at a recirculation rate sufficient tocause intimate intermixing of the residual reaction mixture, the addedfeedstock and the catalyst composition. The heat of the polymerizationreaction is removed from the recirculating intimately intermixedreaction admixture at a rate calculated to provide a substantiallyconstant reaction temperature therein while the same is recirculating inthe reaction zone. The conditions in the reactor are appropriate forcausing olefinic components introduced in the feedstock to undergopolymerization to form the desired polyolefin product in the presence ofthe catalyst composition. A crude product stream containing the desiredpolyolefin product, unreacted olefins and residual catalyst compositionis withdrawn from the reaction zone. The introduction of the feedstockinto the reaction zone and the withdrawal of the product stream from thereaction zone are controlled such that the residence time of theolefinic components undergoing polymerization in the reaction zone isappropriate for production of the desired polyolefin product.

U.S. Patent publication 2003-0040587 A1 published on Feb. 27, 2003 andentitled “Mid-Range Vinylidene Content Polyisobutylene Polymer ProductAnd Process For Producing The Same” (hereinafter the '587 publication)describes a mid-range vinylidene content PIB polymer product and aprocess for making the same. In accordance with the disclosure of the'587 publication, at least about 90% of the PIB molecules present in theproduct comprise alpha or beta position isomers. The alpha (vinylidene)isomer content of the product may range from 20% to 70% thereof, and thecontent of tetra-substituted internal double bonds is very low,preferably less than about 5% and ideally less than about 1-2%. Themid-range vinylidene content PIB polymer products are desirably preparedby a liquid phase polymerization process conducted in a loop reactorsimilar to the reactors described in the '790 application and the '587patent at a temperature which desirably may be about 60° F. or higherusing a BF₃/methanol catalyst complex and a contact time of no more thanabout 4 minutes.

The '790 application, the '587 publication and the '149 patent are eachassigned to the assignee of the present application, and the entiretiesof the respective disclosures thereof are specifically incorporatedherein by this reference thereto.

In conducting the reactions described above, highly specializedequipment may often be employed to enhance the operation and control ofthe polymerization reactors. In each case, for example, the crudeproduct leaving the reactor may be contaminated with residual catalystwhich desirably should be quickly quenched or killed to avoid furtherpolymerization of monomers and low molecular weight oligomers withoutappropriate cooling and/or isomerization resulting from shifting of theposition of the remaining double bond. The catalyst composition may besubjected to contamination by residual materials recirculating with thereaction admixture during the conduct of the polymerization reaction.Moreover, as in any industrial activity, methodology and/or equipmentfor enhancing capacity and throughput are sought continually.

SUMMARY OF THE INVENTION

It is an important aim of the present invention to satisfy the needsdiscussed above. In this regard, in one very important aspect of theinvention, the same provides apparatus for olefin polymerization whichincludes a plurality of reactors. In accordance with the concepts andprinciples of the invention, each of these reactors desirably maycomprise structure defining a reaction zone, an olefin polymerizationreaction mixture inlet connection and a olefin polymerization reactionmixture outlet connection. These connections desirably are in fluidcommunication with the reaction zone. The reactors are each adapted andarranged to facilitate the conduct of an exothermic olefinpolymerization reaction in the reaction zone.

In further accordance with the concepts and principles of the invention,each of the reactors also may include a recirculation system including apump arranged and adapted to circulate the reaction mixture in thereaction zone independently of the introduction of olefin containingfeedstock into the reactor.

The apparatus of the invention also desirably includes an olefincontaining feedstock distribution assembly that comprises an olefincontaining feedstock inlet and a plurality of olefin containingfeedstock outlets. The arrangement of the distribution assembly beingsuch that each of the feedstock outlets is connected in fluidcommunication with the reaction zone of a respective reactor. Theapparatus of the invention may also desirably include a productcollection assembly including a plurality of crude polyolefin productinlets and a crude polyolefin product outlet, the arrangement of thecollection assembly being such that each of the crude polyolefin productinlets is connected in fluid communication with the reaction zone of arespective reactor.

Broadly, the apparatus of the invention may include two or more of thereactors, for example three or four or five or six or more of thereactors.

In another important aspect of the invention, the same provides a methodfor olefin polymerization. In accordance with the invention, the methodincludes providing a plurality of reactors, each of which defines aninternal reaction zone. The method also includes supplying an olefincontaining feedstock, dividing such feedstock into a plurality (2, 3, 4,5, 6 or more) of separate feedstock streams, introducing each of thefeedstock streams into a reaction mixture circulating in the reactionzone of a respective one of the reactors, and conducting an exothermicolefin polymerization reaction in each of the reaction zones.

The method of this aspect of the invention also includes the steps ofseparately circulating the reaction mixture in each reactorindependently of the introduction of the respective stream of feedstockinto the reaction mixture, removing a respective crude polyolefinproduct stream from the reaction mixture circulating in each of thereactors, and combining the crude polyolefin product streams to form asingle crude product stream.

In another aspect, the invention provides a reactor apparatus for olefinpolymerization which comprises at least one reactor defining a reactionzone and including an olefin polymerization reaction mixture inletconnection and an olefin polymerization reaction mixture outletconnection. These connections may desirably be in fluid communicationwith the reaction zone. The reactor is adapted and arranged tofacilitate the conduct in the reaction zone of an exothermic olefinpolymerization reaction on the reaction mixture in the presence of acatalyst composition comprising a catalyst and a catalyst modifier. Inaccordance with this aspect of the invention, the reactor apparatusfurther includes a feedstock inlet, a crude product outlet and arecirculation system including a pump arranged and adapted to circulatethe reaction mixture in the zone independently of the introduction offeedstock into the reaction mixture via said feedstock inlet. Thereactor apparatus of this aspect also includes a catalyst compositioninlet in fluid communication with the zone facilitating the addition ofcatalyst composition to the olefin polymerization reaction mixture andat least one catalyst modifier inlet in fluid communication with thezone facilitating the addition of catalyst modifier to the olefinpolymerization reaction mixture at a rate that is independent of therate of addition of the catalyst composition.

Another important feature of the invention includes the provision of amethod for operating an olefin polymerization reactor. This methodincludes the steps of providing an olefin polymerization reactor havinga reaction zone, recirculating an olefin polymerization reaction mixturein the zone, introducing an olefin containing feedstock into saidreaction mixture, said polymerization reaction mixture beingrecirculated at a flow rate which is independent of the rate ofintroduction of the feedstock into the recirculating olefinpolymerization reaction mixture, introducing a catalyst compositioncomprising a catalyst and a catalyst modifier into the recirculatingolefin polymerization reaction mixture, subjecting the polymerizationreaction mixture to exothermic olefin polymerization reaction conditionsin the zone in the presence of the catalyst composition, and introducinga catalyst modifier into the recirculating olefin polymerizationreaction mixture at a rate that is independent of the rate ofintroduction of the catalyst composition.

In accordance with the concepts and principles of the invention, theforegoing system and methodology may be used in connection with a systemand/or methodology which includes only a single reactor vessel or withone which includes a plurality of reactor vessels arranged in parallelas described above. In this regard, it is to be noted that in accordancewith the invention, the invention further provides an apparatus and/or amethod which includes a multi-reactor system as described above incombination with the described system for introducing catalyst modifierinto the recirculating reaction mixture at a rate that is independent ofthe rate of introduction of the catalyst composition.

It is an additional aspect of the invention to provide a crudepolyolefin product catalyst removal and wash system. In accordance withthis aspect of the invention, the catalyst removal and wash systemcomprises an upstream settler vessel defining an internal settlementchamber adapted and arranged for receiving a mixture of a crudepolyolefin product and an aqueous wash media and allowing the productand the media to separate therein under the influence of gravity. Thesystem further includes a crude, catalyst containing olefinpolymerization product inlet line in fluid communication with thechamber of the upstream settler vessel, a catalyst killing agent inletconduit in fluid communication with the chamber of the upstream settlervessel, and a first make-up water inlet passageway in fluidcommunication with the chamber of the upstream settler vessel.

In addition to the foregoing, the catalyst removal and wash systemdesirably includes a downstream settler system including at least onedownstream settler vessel defining an internal settlement chamberadapted and arranged for receiving a mixture of a partially washed crudepolyolefin product and an aqueous wash media and allowing the productand the media to separate therein under the influence of gravity, anoverhead, partially washed polyolefin product line intercommunicatingthe chamber of the upstream settler vessel with the downstream settlersystem, a washed crude olefin polymerization product outlet line influid communication with the downstream settler system and a secondmake-up water inlet passageway in fluid communication with thedownstream settler system. Finally, the system includes a first drainline intercommunicating the chamber of the upstream settler vessel withan inlet connection to a waste water receiving system, and a seconddrain line intercommunicating the downstream settler system with theinlet connection to the waste water receiving system. Thus, used washwater from the upstream and downstream portions of the system may bepurged separately from the system.

In accordance with the foregoing aspect of the invention, the downstreamsettler system may include one or two or three or more separate settlervessels.

Yet another important feature of the invention is the provision of amethod for treating a catalytically formed crude polyolefin productcontaining residual catalyst to avoid further reaction in the productand remove residual catalyst therefrom. In accordance with this aspectof the invention, the method comprises intimately admixing cruderesidual catalyst containing polyolefin product and a first aqueousmedia containing a catalyst killing agent to thereby form a firstintimately admixed two phase, gravity separable mixture, introducing thefirst two phase mixture into a first settlement zone and allowing thesame to settle in the first zone under the influence of gravity topresent an upper partially washed crude polyolefin product phase and afirst lower aqueous phase containing dissolved catalyst salts,withdrawing the first lower aqueous phase from the first settlement zoneand recirculating a first portion thereof and introducing the same intothe first two phase mixture for inclusion as part of the first aqueousmedia, directing a second portion of the first lower aqueous phase to adrain for disposal or reclamation, introducing a first quantity ofmake-up water into the first two phase mixture for inclusion as part ofthe first aqueous media, withdrawing the partially washed crudepolyolefin product phase from the first settlement zone and intimatelyadmixing the same with a second aqueous media to thereby form a secondintimately admixed two phase, gravity separable mixture, introducing thesecond two phase admixture into a second settlement zone and allowingthe same to settle in the second zone under the influence of gravity topresent an upper more fully washed crude polyolefin product phase and asecond lower aqueous phase, withdrawing the second lower aqueous phasefrom the second settlement zone and recirculating a first portionthereof and introducing the same into the second two phase mixture forinclusion as part of the second aqueous media, directing a secondportion of the second lower aqueous phase to a drain for disposal orreclamation, removing the more fully washed crude polyolefin productphase from the second settlement zone, and introducing a second separatequantity of make-up water into the second two phase mixture forinclusion as part of the second aqueous media.

It is to be noted that the catalyst removal and wash system and/ormethod described above is suitable for use in connection with either asystem which includes only a single reactor as well as one whichincludes a plurality of reactors as described above. Thus, it is animportant aspect of the invention to provide a system and/or methodwhich includes both a multi reactor system and a catalyst removal andwash system and/or method as described. In addition, such combinedsystem may also include the described system for adjusting the amount ofcatalyst modifier in the recirculating reaction mixture.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 is a schematic illustration of a reactor system including amulti-pass shell and tube heat exchanger and a recirculation systemwhich is useful in connection with the invention;

FIG. 2 is a flow diagram illustrating an apparatus which embodies theconcepts and principles of the invention and which employs two reactorsof the sort illustrated in FIG. 1 arranged for operation in parallel;and

FIG. 3 is a flow diagram illustrating a system for receiving a crudepolyolefin product from the apparatuses of FIGS. 1 and 2, for example,and treating the same to wash the crude product and remove residualcatalyst therefrom.

DETAILED DESCRIPTION OF THE PREFERRED EMBODIMENTS

Many potentially valuable reactors having utility in the conduct ofliquid phase polymerization polyolefins are known to the routineers inthe art to which the invention pertains. However, for purposes of onepreferred embodiment of the present invention, the reactor desirably mayinclude a two-pass shell-and-tube heat exchanger as shown in FIG. 1,where the same is identified by the numeral 10. The reactor 10 may, forexample, include three hundred eighty eight (388) 0.375″ tubes with awall thickness of 0.035″, each thereby providing an internal tubediameter of 0.305″. The reactor may be twelve feet long and may haveinternal baffling and partitions to provide 2 passes with 194 tubes perpass. The passes are identified by the numerals 50 and 51 in FIG. 1, andthe 194 tubes of each pass are respectively represented by the singletube portions 52 and 53. Such construction is well known in the heatexchanger and reactor arts and no further explanation is believednecessary.

In operation, an olefin (e.g., isobutylene, 1-butene, 2-butene)containing feedstock enters the reactor system via pump 14 and pipe 15.The downstream end of pipe 15 desirably may be located to direct thefeed stock into the suction line 20 of recirculation pump 25. A catalystcomposition may be injected into the reactor circulation system via pump29 and pipe 30 at a location downstream from pump 25 and adjacent thefirst pass as shown in FIG. 1. The catalyst composition may desirably bea methanol/BF₃ complex with a molar ratio of methanol to BF₃ of about1.9:1 or less and preferably may be a methanol/BF₃ complex with a molarratio of methanol to BF₃ of about 1.7:1 or less. Desirably the molarratio of methanol to BF₃ may be as low as about 1.1:1 or less for someapplications.

Circulation pump 25 pushes the reaction mixture through line 35, controlvalve 40 and line 45 into the bottom head 11 of the reactor 10. A flowmeter 46 may be positioned in line 45 as shown. Appropriate temperatureindicators TI and pressure indicators PI may be provided to monitor thesystem. The reaction mixture travels upwardly through tubes 52 of pass50 and downwardly through tubes 53 of pass 51. The circulating reactionmixture leaves reactor 10 via suction line 20. The reactor system thusis of the type which is sometimes referred to as a loop reactor. Withthis system, which is only a preferred system since there are many otherarrangements which would be apparent to those of ordinary skill in theart, the flow rate of the reactant mixture in the reactor may beadjusted and optimized independently of feed stock introduction andproduct removal rates so as to achieve thorough intermixing of thecatalyst composition and the reactants and appropriate temperaturecontrol.

As explained previously, each pass 50 and 51 may desirably include onehundred ninety four (194) separate tubes. For clarity, however, only aportion of a single tube is illustrated schematically in each pass inFIG. 11. These tubes are identified by the respective reference numerals52 and 53. Although only a portion of each representative tube 52 and 53is shown, it should be appreciated by those skilled in the art that eachof these tubes extends for the entire distance between top head 12 andbottom head 11 and that the same are in fluid communication with theinteriors of heads 11 and 12.

It is to be noted here, that the reaction mixture should preferably becirculated through the tubes 52, 53 of the reactor at a flow ratesufficient to cause turbulent flow, whereby to achieve intimateintermixing between the catalyst composition and the reactants and aheat transfer coefficient appropriate to provide proper cooling. In thisregard, the flow rate, the reaction mixture properties, the reactionconditions and the reactor configuration should be appropriate toproduce a Reynolds number (Re) in the range of from about 2000 to about3000 and a heat transfer coefficient (U) in the range of from about 50to about 150 Btu/min ft²° F. in the tubes 52, 53 of the reactor. Suchparameters may generally be obtained when the linear flow rate of atypical reaction mixture through a tube having an internal diameter of0.331 inch is approximately within the range of from about 6 to 9 feetper second.

A product exit line 55 may desirably be connected in fluid communicationwith pump suction line 20. However, as would be readily appreciated bythose skilled in the art, the exit line could be positioned almostanywhere in the system since, at least from a theoretical view point,and as explained below, the conditions in the reactor may desirablyapproach those of a continuous stirred tank reactor (CSTR) where bothtemperature and composition remain constant such that the composition ofthe product stream leaving the reactor is identical to the compositionof the reaction mixture recirculating in the reactor. Likewise, thefeedstock introduction line 15 could be positioned almost anywhere inthe system, although, in practice, it is desirable for the line 15 to beconnected into the recirculation system at a position that is as farupstream from the line 55 as possible to insure that monomers introducevia line 15 have a maximum opportunity to polymerize before encounteringline 55.

A coolant may desirably be circulated on the shell side of the reactorat a rate to remove heat of reaction and maintain a desired temperaturein the reactor.

A catalyst complexing agent may desirably be added to the circulatingreaction mixture via pump 18 and line 16 positioned in top head 12. Thisfeature is particularly valuable when the desired product is highlyreactive polyisobutylene (HR PIB) and the catalyst composition comprisesa BF₃ catalyst and a methanol complexing agent. BF₃ complexes withmethanol in two different forms, i.e., mono-complex (1 mole of BF₃ to 1mole of methanol) and di-complex (1 mole of BF₃ to 2 moles of methanol),depending upon the availability of methanol. The mono-complex is thetrue catalytic species, whereas the di-complex does not have anyparticular catalytic properties in the absence of the mono-complex.References to fractional complexes are the actual average of themono-complex and the di-complex. In this regard it has been determinedthat a catalyst composition made up of 0.59 to 0.62 moles of BF₃ permole of methanol is particularly valuable in the production of HR PIB.But when such a composition is introduced into the system, variationsand contaminants in the hydrocarbon feed often may result in less thanoptimal reactor control. This is believed to be, at least in part, theresult of the propensity for many contaminants to effectively increasethe apparent ratio of methanol to catalyst in the composition. Moreover,it is not always possible to predetermine the exact contamination levelof some feedstocks.

In accordance with the concepts and principles of the invention,however, it has been discovered that these problems may be solved andoptimal results may be achieved simply by introducing a catalystcomposition, which for some purposes may desirably be a methanol leancomposition, e.g., one containing more than the optimum desiredconcentration of the mono-complex, into the reactor 10 via line 30, andindependently adding relatively pure methanol through a line that maydesirably be spaced from line 30, such as the line 16. A pump 18 maydesirably be provided to push the methanol through pipe 16.Alternatively, essentially the same effect may be achieved byintroducing a separate methanol stream directly into the catalystcomposition stream in line 30 by way of a line (not shown) andintroducing the added methanol and the catalyst composition into thesystem together. In either event, the additional methanol is availableto trim the catalyst composition so that a desired methanol to BF₃ ratiomay be achieved and maintained in the reactor 10.

In further accordance with the concepts and principles of the invention,the amount of methanol added should desirably be sufficient to createand maintain a preferred ratio of BF₃ per mole of methanol in thecirculating reaction mixture. For some applications, for example where ahighly reactive polyisobutylene is the desired product, the catalystcomposition added via line 30 may desirably comprise a molar ratio ofBF₃ and methanol in the range of from about of 0.59:1 to about 0.62:1,and ideally may be about 0.61:1. Alternatively, for other applications,for example where the desired product is a polyisobutylene where thevinylidene content is not so important, the catalyst composition addedvia line 30 may ideally comprise a molar ratio of BF₃ and methanol ofabout 1:1.

The product exiting the system via line 55 should be quickly quenchedwith a material capable of killing the activity of the catalyst, suchas, for example, ammonium hydroxide, so that the ongoing exothermicpolymerization reactions are immediately stopped. Thus, any undesiredtemperature increase due to a lack of cooling (and the concomitantproduction of lower molecular weight polymers due to the highertemperatures) or rearrangement of the polymer molecules may beminimized. The polyolefin products of the invention may then be directedto a work up system, including a wash system as described below, wherecatalyst salts may be removed and a purification and separation system(not shown) where the polyolefin product may be separated from unreactedmonomers, dimers, oligomers and other undesirable contaminants such asdiluents, etc. These latter materials may then be recycled or divertedfor other uses employing known methodology.

With the described recirculation system, the rate of feedstockintroduction into the reaction mixture and the rate of product removalare each independent of the circulation rate. As will be appreciated bythose of ordinary skill in the art, the number of passes through thereactor and the size and configuration of the latter are simply mattersof choice. For a single reactor system as illustrated in FIG. 1, thefeedstock and product withdrawal flow rates may preferably be chosensuch that the residence time of the fresh monomers entering the reactorwith the feedstock is 4 minutes or less, desirably 3 minutes or less,preferably 2 minutes or less, even more preferably 1 minute or less, andideally less than 1 minute. In connection with the foregoing, theresidence time is defined as the total reactor system volume divided bythe volumetric flow rate of the feedstock entering the system via pipe15.

The recirculation flow rate, that is the flow rate of the reactionmixture in the system induced by the recirculation pump 25, iscontrolled, as described above, to achieve appropriate turbulence and/orheat transfer characteristics. This recirculation flow rate is often afunction of the system itself and other desired process conditions. Forthe systems described above, the ratio of the recirculation flow rate tothe incoming feedstock flow rate (recycle ratio) should generally bemaintained in the range of from about 20:1 to about 50:1, desirably inthe range of from about 25:1 to about 40:1, and ideally in the range offrom about 28:1 to about 35:1. In particular, in addition to causingturbulence and providing an appropriate heat transfer coefficient, therecirculation flow rate of the reaction mixture should be sufficient tokeep the concentrations of the ingredients therein essentially constantand/or to minimize temperature gradients within the circulating reactionmixture, whereby essentially isothermal conditions are established andmaintained in the reactor.

As mentioned above, the recycle ratios generally should be in the rangeof from about 20:1 to about 50:1. Higher recycle ratios increase thedegree of mixing and the reactor approaches isothermal operation leadingto narrower polymer distributions. But higher recycle ratios also resultin higher power consumption. Lower recycle ratios decrease the amount ofmixing in the reactor, and as a result, there is a greater discrepancyin the temperature profiles. As the recycle ratio approaches zero, thedesign equations for the reactor reduce to those for a plug flow reactormodel. On the other hand, as the recycle ratio approaches infinity, themodeling equations reduce to those for a CSTR. When CSTR conditions areachieved, both temperature and composition remain constant and thecomposition of the product stream leaving the reactor is identical tothe composition of the reaction mixture recirculating in the reactor.Needless to say, after equilibrium has been established, as feedstockenters the system, an equal mass of product is pushed out of the reactorloop. Thus, under CSTR conditions, the point at which the product streamis withdrawn is independent of reactor geometry.

The feedstock entering the system through line 15 may be any olefincontaining stream. Where polyisobutylene is the preferred product, thefeedstock may be, e.g., isobutylene concentrate, dehydro effluent, or atypical raff-1 stream. These feedstock materials are describedrespectively below in Tables 1, 2 and 3.

TABLE 1 Isobutylene Concentrate Ingredient Weight % C₃s 0.00 I-butane6.41 n-butane 1.68 l-butene 1.30 I-butene 89.19 trans-2-butene 0.83cis-2-butene 0.38 1,3-butadiene 0.21

TABLE 2 Dehydro Effluent Ingredient Weight % C₃s 0.38 I-butane 43.07n-butane 1.29 l-butene 0.81 I-butene 52.58 trans-2-butene 0.98cis-2-butene 0.69 1,3-butadiene 0.20

TABLE 3 Raff-1 Ingredient Weight % C₃s 0.57 I-butane 4.42 n-butane 16.15l-butene 37.22 I-butene 30.01 trans-2-butene 8.38 cis-2-butene 2.271,3-butadiene 0.37 MTBE 0.61

On the other hand, suitable streams for the production of polyolefinsgenerally include feedstock materials such as those described in Tables4 and 5.

TABLE 4 2-Butene Rich Stream Ingredient Weight % I-butane 2.19 n-butane61.50 l-butene 0.64 trans-2-butene 28.18 cis-2-butene 7.49

TABLE 5 1-Decene Rich Stream Ingredient Weight % 1-decene 94.00 C₁₀isomers 6.00

With reference to FIG. 2, and in further accordance with the conceptsand principles of the invention, it has been discovered unexpectedlythat an operating system incorporating a plurality of reactors arrangedfor operation in parallel provides a great deal more operatingflexibility than a single larger reactor sized for the same totalproduction rate. In fact, the multiple reactor concept of the inventionprovides for less risk in operation, more flexibility in running theprocess, lower feed rates (higher conversions), improved reactor design,and increased production capability per unit of time. Moreover, themultiple reactor concept of the invention allows, e.g., for a 20:1scale-up from pilot plant operation when the system includes tworeactors, rather than a 40:1 scale-up with a larger reactor. Thissignificantly reduces the uncertainties associated with scaling up pilotplant data.

A multiple reactor system which embodies the concepts and principles ofthe invention is illustrated in FIG. 2, where it is identified broadlyby the reference numeral 200. System 200 includes two reactors 202 a and202 b, which as shown are connected for parallel operation on both thereaction side and on the cooling fluid side. In addition, each reactor202 a, 202 b, desirably has its own respective recirculation system, 204a, 204 b. Ideally, the reactors 202 a and 202 b may be identical.However, in accordance with the broad aspects of the invention, it isnot a critical feature of the invention for the reactors to beidentical.

Ideally, the reactors 202 a, 202 b may each be essentially the same asthe reactor 100 illustrated in FIG. 1. That is to say, the reactors 202a, 202 b may each be a two-pass reactor, with each pass including onehundred ninety four ⅜″ tubes as described above. Other equipment shownin FIG. 2 which is essentially the same as the corresponding equipmentshown in FIG. 1 is identified by similar reference numerals followed byeither an “a” or a “b” as the case may be. Thus, the reactors 202 a, 202b each include a feedstock inlet line (15 a, 15 b), a recirculation pump(25 a, 25 b), a recirculation pump suction line (20 a, 20 b), a productoutlet line (55 a, 55 b), a catalyst composition inlet line (30 a, 30 b)and a methanol inlet line (16 a, 16 b). In FIG. 2, a common feedstockinlet line for the multiple reactor system 200 is identified by thereference numeral 215, and a common product outlet line for the multiplereactor system 200 is identified by the reference numeral 255.

The multiple reactor system of the invention offers advantages inconversion and polymer polydispersity. The multiple reactor system ofthe invention also facilitates a reduction in the amount of off-specmaterial generated during early operation of the unit becauseequilibrium and the development of the operating parameters necessaryfor a particular product are achieved more expeditiously.

The optimum inlet feedstock flow rate for each reactor of the multiplereactor system 200 of the invention is about fifteen to seventeengal/min with appropriate refrigeration capacity and back-end processingcapabilities. That is to say, with the multiple reactor system 200 ofthe invention, higher conversions (70-75%) are possible at this flowrate than higher flow rates (>20 gal/min per reactor). This is theresult of increased residence times in the range of from about 120 to135 seconds. Higher conversion rates lead to improvements (reductions)in polydispersity, and a polydispersity of about 1.7 is achievablethrough the use of the multiple reactor system 200 of the invention forthe production of a PIB product having a number average molecular weight(M_(N)) of about 950, and a polydispersity of about 2.2 is achievablethrough the use of the multiple reactor system 200 of the invention forthe production of a PIB product having a M_(N) of about 2300. When usinga single reactor to produce the same molecular weight products, the bestpolydispersities that could be achieved were 1.9 and 2.3 respectively.

For the dual reactor system described above, the feedstock and productwithdrawal flow rates may preferably be chosen such that the residencetime of the reaction mixture within each reactor may be, for example,about 4 minute or less, about 3 minutes or less, ideally from about 120to about 135 seconds, perhaps even less than about 2 minutes, andpotentially even as low as about 1 minute or less.

The multiple reactor system 200 of the invention also facilitates theuse of smaller reactors having improved pressure drop characteristicsresulting in more efficient energy usage. This may be due at least inpart to the fact that larger reactors may require longer reactor tubeswith similar recirculation linear flow rates.

EXAMPLE

Tests were conducted to determine the improvements in operationalcharacteristics achievable through the use of a multiple reactor system,in this case using two similar reactors operating in parallel. Accordingto the test protocol, the tests were conducted in three phases. In thesephases, all operating parameters other than those specifically spelledout were held constant. In the first phase, a single reactor wasoperated in a manner to produce a highly reactive polyisobutylene havingterminal double bond content greater than 70% and a M_(N) ofapproximately 1600. The feedstock was an isobutylene concentrate and therecirculation rate was maintained at a level to achieve intimateintermixing between the catalyst composition and the reactants and aheat transfer coefficient appropriate to provide proper cooling. Thesingle reactor was initially operated with a feedstock inlet rate of 27gpm. Later, this was increased to 32 gpm. In the second phase, tworeactors were operated in parallel. These parallel reactors were eachessentially the same as the reactor employed during the first phase.During this phase, the feedstock inlet rate to each reactor was 15 gpm.And again, the recirculation rate was maintained at a level to achieveintimate intermixing between the catalyst composition and the reactantsand a heat transfer coefficient appropriate to provide proper cooling.In the third phase, the setup was the same as in the second phase. As aninitial step in this third phase, the conversion rate was increasedwhile the feedstock inlet rate to each reactor was maintained at 15 gpm,then the chilled water supply to the shell side of the reactors wasreduced to increase the conversion rate. Thereafter, the feedstock inletrate to each reactor was increased to 17 gpm. The results of these testsare summarized below in Table 6.

TABLE 6 Summary of Rate Test Results Hydrocarbon Heat ReactorRefrigeration Data Test Feed Reactor Balance Make Flow Supply RequiredPhase^(a) Test Duration Reactor Flow (gpm) Temp. (° F.) Conv. (%) Rate(lb/min) (gpm) Temp (° F.) Tons 1 22 hrs. A 31.7 70 62 89.6 593 29.4 1642 45 hrs. A 15 64.5 67 45.9 257 38.8  86 B 15 64.5 67 45.9 252 38.8  8730 91.8 509 173 3.1 22 hrs and 40 min. A 15 64 73 49.9 244 29.8  95 B 1564 73 49.9 249 29.8  95 30 99.8 493 189 3.2 3 hrs and 10 min. A 17 64 6852.6 244 31.3  99 B 17 64 68 52.6 249 31.3  99 34 105.2 493 198^(a)Phase 1 — One reactor, increased feed rate to maximize reactor makerate ^(a)Phase 2 — Two reactors at baseline feed rates of 15 gpm, 38.8°F. chilled water ^(a)Phase 3.1 — Two reactors at 15 gpm decreasedchilled water temperature to maximize conversion ^(a)Phase 3.2 — Tworeactors, increased feed rate to maximize reactor make rate

Table 6 presents the length of each test phase, feed flow rate, reactiontemperature, heat balance conversion, reactor make rate, andrefrigeration system data. The heat balance conversion was estimatedbased on the feed flow, heat of reaction, and the chilled water flow andtemperature increase across the reactor. The chilled water flow andtemperature increase determine the amount of heat generated by thereaction, and the heat of reaction and feed flow are used to calculatethe percentage of the feed converted to PIB and oligomers. The reactormake rate, in pounds per minute (lb/min), is calculated from the feedrate and heat balance conversion.

During single reactor operation, the reactor make rate was maximized at88 lb/min with the feed rate at 31.7 gpm. When the feed rate wasincreased to 32 gpm, the reactor make rate began to drop, so the feedrate was not increased any further. The highest reactor make rate wasachieved during two reactor operation at a feed rate to each reactor of17 gpm. While at 15 gpm feed to each reactor, the reactor make rate wasincreased from 45.9 lb/min per reactor (91.8 lb/min total) to 49.9lb/min per reactor (99.8 lb/min total) by reducing the chilled watersupply temperature from 38° F. to 30° F. The reactor make rate wasincreased further to 52.6 lb/min per reactor (105.2 lb/min total) byincreasing the feed rates to each reactor from 15 gpm to 17 gpm.

During Phase 1 of the test program, the feed rate was held at 30.5 gpmfor about 8 hours. A direct comparison can be made between one and tworeactor operation by comparing the conversion during this period withthe conversion obtained during Phase 3.1. During Phase 1, the feed ratewas slightly higher (30.5 versus 30 gpm), but the chilled water supplytemperature was slightly lower (28° F. versus 30° F. during Phase 3.1).With two-reactor operation, the heat balance conversion was 73% versus64% for one reactor operation, even though the reactor temperatures wereoperated 5° F. cooler when operating two reactors (64° F. versus 69°F.). With two-reactor operation, there is twice the residence time andtwice the surface area for removing heat compared to the single reactorcase. The additional residence time explains why the reactor temperaturehad to be lowered, and the additional surface area explains why theconversion was higher even at the lower reaction temperature.

In view of the foregoing, it can readily be seen that when two reactorsare used in parallel, conversion rates are increased and polydispersityis lowered relative to a single reactor. This result is achieved becausethe multiple reactor concept facilitates a lower feed rate to eachreactor with a concomitant increase in the residence time.

As mentioned above, product exiting the polymerization reactor systemvia lines 55 (FIG. 1) or 255 (FIG. 2) should be quenched immediatelywith a material capable of killing the activity of the catalyst, suchas, for example, ammonium hydroxide. Thus, any potential undesireddecrease in molecular weight or rearrangement of the polymer moleculemay be minimized. The polyolefin products of the invention may then bedirected to a work up system, including a wash system as described belowwhere catalyst salts may be removed.

In FIG. 3, a wash system which embodies the concepts and principles ofanother aspect of the invention is identified broadly by the referencenumeral 300. As shown, the system 300 includes an upstream settlervessel 302 and a downstream settler system 304 which, in the preferredembodiment of this aspect of the invention shown in FIG. 3, includes twodownstream settler vessels 306, 308. It is to be noted here, that in thealternative, the downstreamn settler system 304 could just as wellinclude only a single settler vessel or three or more settler vessels,depending upon the nature of the product and the nature of the residualcatalyst materials to be removed therefrom.

System 300 further includes an inlet line 310 which interconnects eitherline 55 or line 255, as the case may be, and the suction 311 of a pump312 which pumps crude product and materials intermixed therewith intosettler vessel 302 via line 314. An agent for killing the activity ofany residual catalyst in the crude product entering system 300 via line310 is introduce into line 310 via pump 316 and line 318. NH₄OH in anaqueous solution is a particularly good agent for killing the activityof any residual BF3/methanol complex in the polyolefin product. However,the invention is in no way limited to the use of NH₄OH. Rather, theexact nature of the catalyst activity killing agent will depend entirelyupon the nature of the catalyst itself and/or the nature of the productin the product stream.

Wash water is introduced into and admixed with the crude product in line316 via a line 320. The admixture of crude product containing residualcatalyst composition, the catalyst activity killing agent and the washwater is introduced into the pump 312 via suction line 311. Desirably,the pump 312 may be centrifugal pump wherein the rotation of theimpellers insures that the water, catalyst salt resulting from theinteraction between the catalyst activity killing agent and the catalystand the crude polyolefin product are intimately intermixed such thatthorough washing is achieved. In addition, pump 312 may be provided witha recycle line 322, including a flow controlling device 324, to return aportion of the admixture to the pump suction for additional mixing.

The admixture of hydrocarbon product, killed catalyst salts and waterare introduced via line 314 into an internal settlement chamber of thesettlement vessel 302 where the hydrocarbon phase is separated from theaqueous phase under the influence of gravitational forces in a mannerthat is known per se. Desirably, in this latter regard, the interactionbetween the catalyst activity killing agent and the catalyst forms awater-soluble salt such that the bulk of such salt will be present inthe aqueous phase.

The upper, partially washed crude polyolefin product phase is removedfrom vessel 302 via an overhead line 326 and the aqueous phase leavesvessel 302 via a line 328. A portion of the removed aqueous phase isrecycled to the wash water inlet line 320 via return line 330 and a flowcontroller 332. Another portion of the removed aqueous phase isdiscarded from the system via a drain line 334 and a level controller336 which controls the level of the aqueous phase in vessel 302. Drainline 334 is connected to a system (not shown) for either reclamation ordisposal of the used and contaminated wash water.

Make-up wash water for vessel 302, which desirably may be demineralizedwater, is added to the recycled drain water in line 320 via a line 321.In this connection it is to be noted that the respective amounts ofmake-up wash water, incoming catalyst killing agent, and purged aqueousphase should all be controlled so as to insure that the amount ofkilling agent entering the system is always in an excess relative to theamount of residual catalyst in the crude product.

The partially washed crude polyolefin product phase in line 326 isintroduced into the suction 336 of a pump 338 along with additional washwater delivered via line 340. Pump 338 may desirably be a centrifugalpump like pump 312 to ensure intimate admixing between the water phaseand the hydrocarbon phase before the admixture is delivered via line 342into vessel 306. Pump 338 may also be equipped with a recycle line 344and a flow controlling device 346 to return a portion of the admixtureto the pump suction 336 for additional mixing. The two phase admixturein vessel 306 is allowed to separate under the influence of gravity toform an upper hydrocarbon phase and a lower aqueous phase.

An upper, more thoroughly washed crude polyolefin product phase isremoved from vessel 306 via another overhead line 348, and the lowersettled aqueous phase leaves vessel 306 via a line 350. A portion of theremoved aqueous phase is recycled to wash water inlet line 340 viareturn line 352 and another portion of the removed aqueous phase ispurged from the system via a drain line 354 and a level controller 356which controls the level of the aqueous phase in vessel 306. Drain line354 is connected to drain line 334.

The more thoroughly washed crude polyolefin product phase in line 348 isadmixed with additional wash water which is introduced via line 358. Theadmixture of crude polyolefin product phase and additional wash water isintroduced into settler vessel 308 via a line 359 where once again thetwo phase admixture is allow to separate under the influence of gravity.The lower aqueous phase is removed from vessel 308 under the influenceof a pump 362 via a lower line 360, and a portion thereof is recycled toline 348 via a flow controller device 364, a return line 366 and line358. Another portion of the aqueous phase leaving vessel 308 is recycledvia line 367 and flow controller 369 and introduced into line 340 foruse as make-up wash water in vessel 306.

The completely washed crude polyolefin product is removed from vessel308 via an overhead line 374 and forwarded to a downstream purificationsystem (not shown) for the removal of diluents, unreacted monomer, andunwanted light ends such as dimers, trimers, oligomers, etc.

Fresh make-up wash water for vessel 308, which once again may desirablybe demineralized water, is introduced via line 368. The make-up water isintroduced into the system via a pump 370 and line 372. In this regardit is to be noted that line 372 is connected with line 321 to providefresh make-up water for upstream vessel 302 and with line 368 toseparately and independently provide fresh make-up water for downstreamsettler system 304. Accordingly, extra make-up water may be introducedinto the downstream settler system 304 without unnecessarily dilutingthe catalyst activity killing agent (NH₄OH) needed in the upstreamsettler vessel 302. This result is achieved because the wash system forthe upstream vessel is operated completely independently of the washsystem for the downstream settler system 304. It is also noteworthy thatthe concentration of the catalyst activity killing agent in the aqueousphase of the upstream settler vessel 302 should always be in excessrelative to the amount of residual catalyst. Moreover, the concentrationof the catalyst salts in the aqueous phase should always be low enoughto avoid precipitation. Accordingly, the amount of fresh make-up waterintroduced into the upstream settler vessel 302 needs to be closelycontrolled, while the amount of fresh make-up water introduced into thedownstream settler system should be copious and determined solely by theneed for removing as much contamination from the final product aspossible. Thus, a lower flow of fresh make-up water is used in theupstream settler vessel to minimize the usage of the catalyst activitykilling agent, while a much greater flow of fresh make-up water is usedin the downstream settler system to provide for better washing.

1. A method for olefin polymerization comprising: providing a reactorsystem including a plurality of reactors arranged for parallel flow,each said reactor defining an internal reaction zone; supplying anolefin containing feedstock and dividing the same into a plurality ofseparate feedstock streams; introducing each of said separate olefincontaining feedstock streams into the reaction zone of a respective oneof said reactors; conducting an exothermic olefin polymerizationreaction on an olefin polymerization reaction mixture in each reactionzone; separately circulating the reaction mixture in each reactor at aflow rate that is independent of the rate of introduction of therespective stream of feedstock into the reaction zone while saidexothermic olefin polymerization reaction is ongoing; removing a crudepolyolefin product stream from each of said reactors; and combining saidcrude polyolefin product streams to form a single crude product stream.2. A method as set forth in claim 1, wherein said system includes two ofsaid reactors and said olefin polymerization reaction mixture is dividedinto two separate streams.
 3. A method as set forth in claim 1, whereinsaid system includes at least three of said olefin polymerizationreaction mixture is divided into at least three separate streams.